Optimum process for selective hydrogenation/hydro-isomerization, aromatic saturation, gasoline, kerosene and diesel/distillate desulfurization (HDS). RHT-hydrogenationSM, RHT-HDSSM
A process for selective hydrogenation of C3/C4/C5/C6/C7 and LCN, hydro-isomerization of olefins, benzene saturation and hydrodesulfurization of Gasoline, Kerosene and Diesel together with aromatic saturation of LCO is being provided so as to provide optimum technology at low cost with unique configuration. The technology is user friendly and uses conventional catalyst. These configurations cover both the options of installing the bulk catalyst in the distillation column with chimneys trays, hence essentially all the reaction takes place in the liquid phase in single or multiple beds or with fixed bed down flow or up flow reactor configuration. The configurations shown in the figures depicts that the fixed bed reactors are integrated with the Distillation Column and this art helps in lowering the Capital costs and the reactors are operating in single phase or two-phase conditions. The art is applicable to MAPD, vinyl acetylene, C3, C4, C5, C6, C7 mixed hydrocarbon stream, and LCN diolefin selective hydrogenation. It includes the process for hydro-isomerization of Butene-2 to Butene-1, or visa versa, removal of Isobutylene and Isobutane by distillation after hydro isomerization, benzene hydrogenation to Cyclohexane, hydrodesulfurization of FCC gasoline, coker gasoline or any other Naphtha stream together with hydrodesulfurization of Diesel/Kerosene from any of the refinery units.
The inventions and the arts here in the following processes is selective hydrogenation of Acetylene, MAPD, C3/C4/C5/C6/C7 and LCN, Hydro-isomerization, Benzene saturation and hydrodesulfurization of Gasoline, Kerosene and Diesel/Distillate in a unique and optimized configuration and selection of multiple catalysts so as to provide low cost processes and maximizing yields compared to Conventional processes. RHT process is enhanced configuration compared to process with simultaneous reaction and distillation which uses the high cost catalyst and is cumbersome to load in the equipment, together other potential drawbacks as Reaction conditions are not normally consistent with fractionation which provides a non optimum option. This Reactive distillation technology also requires much higher energy due to reflux requirements for the column. The present art of RHT process provides an alternate to conventional technology and Reactive distillation that is being applied for Selective Hydrogenation, Benzene Saturation and Gasoline HDS applications. In Benzene Saturation and Gasoline HDS applications it is not logical to use the Reactive Distillation as it increases the Capital cost due to high temperature energy source for benzene saturation, which might require a Furnace. In the case of Gasoline HDS, Reactive distillation, requires the reflux for the column, and to enhance the WABT of the catalyst zone (at operating pressure) most of the product is taken overhead, (essentially column working in a recycle mode), which doubles the energy requirements of the process (Licenser's patent U.S. Pat. No. 6,495,030). This is an expensive alternative, when energy costs are 75 $/bbl of oil, though heat integration can improve the energy utilization but still will have to lose lot of the hot energy to waste. RHT technologies apart from providing unique optimized configurations by providing alternate reaction and distillation in the column by installing bulk catalyst in the column or outside in the side reactors, provides one unit operation in the single equipment or alternate integrating fixed bed reactors to the Column as side reactor and are capable of operating in down flow or up flow mode, which ever gives better economics for particular application. Reactor and Distillation column are operated at best operating conditions suitable for each unit operation rather than sacrificing the economics for Reactive distillation on the basis of higher catalyst life, which can be achieved by removing the diolefins from the feed and operating the reactor in two phase flow operation. RHT provides the concept in this art, with reactions and separation done at most favorable conditions and not to install high cost and cumbersome catalyst in the column providing a mirage that the economics is better. The
The technology and the art of RHT provides process for mercaptan removal from FCC Naphtha (
Following typical Catalyst applications are being proposed: for selective hydrogenation, Isomerization, aromatic saturation and Hydrodesulfurization of Gasoline, kerosene and Diesel/distillates. The invention's advantages can be easily seen by people familiar in art and the merits of different catalyst for different applications but are not limited to these catalysts only. The art also allows the user to switch catalyst based on the availability in the market place if some new catalyst comes to the market; the information provided here is just for illustration. The processes/configurations use different catalyst for Selective Hydrogenation and isomerization, hydrodesulfurization, aromatic saturation and also for desulfurization of cracked and straight run Feeds as shown below.
Following catalyst are used for the services proposed in this invention and are not limited to these and can use any commercial available catalyst for the service. The unique and superior configuration provides advantages in selectivity, low cost of catalyst and utilities. Cost of extra equipment is paid up in 6 to 12 months depending on service compared to some technologies which claim lower capital cost. The other advantages of selectivity/Low Octane loss, low energy has been not taken into account for this patent economics as regards to payout period for extra equipment. This essentially means that after taking all the benefits there are major advantages provided by this configuration.
As mentioned above that the catalyst could be bulk catalyst from any of the above catalysts but one is not limited to these catalyst only for all the above processes and would chose what ever best catalyst is available in the market. The bulk catalyst in the column or in the side reactors provides high catalyst efficiency and usage and lower catalyst cost and better yield/selectivity and also low utility consumption. There is not much of an advantage of having the catalyst in packaging (which is usually is cumbersome to install/loading) in the column and simultaneous reaction and distillation, except few processes where rate constants are high but that can also be compensated by unique RHT configuration as explained in this and other patents (U.S. Pat. No. 4,503,265). In most of the applications the reaction and distillation operating conditions are not optimum for each unit operation. By artificial means of increasing the certain parameters to make the process workable in this application provides the solution which are not cost effective due much higher and catalyst and operation being at not optimum conditions. RHT provides the application and configuration, with bulk catalyst where both distillation and reaction can be essentially decoupled and run at optimum conditions, but still integrated with the column so as to save the cost. The bulk catalyst in the column or in side reactors attached to column can be loaded above and below the feed location but preferably above the feed location. In some applications the fixed bed reactor can also be installed upstream of the column as shown in the
The art of this application is applicable to following processes. RHT-HydrogenationSM, RHT-HDSSM
Selective Hydrogenation of Vinyl acetylene, methyl acetylene/propadiene, butadiene, isoprene, pentadienes, hexadiens etc. and other diolefins in hydrocarbon streams from FCC, Steam Cracker, Thermal Crackers (e.g. Visbreaker and delayed Cokers etc.) are reduced by this application to 5 wppm to 1000 wppm in the product as per the requirements of the process. The conditions for Operating the reactor will depend on each species and will have Ni, Pd, Pd/Pt. Pd/Ag, Ni/Mo or Ni/W catalyst but not limited to, and having stoichiometric amount of hydrogen to about 200 Scf/bbl of Feed, preferably around 50 to Scf/bbl of Feed. The temperature range will be in 100 F. to 400 F., but depending on the Feed composition and catalyst type, temperature of around 100 F. to 300 F. is used. The pressure range will be in 75 psig to 400 psig, but preferably in the range of 150 to 300 psig. In most of the cases the fractionator could be operated at lowest pressure so as to condense the overhead product. The pressures/temperatures ranges are provided for the selective hydrogenation reactor applications as shown in the FIGS. 1,2, 3, 4 and 5.
Isomerization can be done with Pd, Pd/Pt, Ni or Ni/Mo catalyst but not limited to these catalysts, based on the Feed characteristics and for most of the applications temperature could be in the range of 100 to 300 F., similar operating conditions required for Selective hydrogenation (SHU), unit as mentioned above. This is shown in
Aromatic Saturation/hydrogenation, this can be done with Ni or Pt/Zeolite catalyst depending on the impurities and ultimate specification. For Ni one need to be careful with Feed impurities, as stated in the literature and well described in the art that sulfur in the feed should be less than 1 wppm as it poisons the catalyst (though reformate product has no sulfur unless contamination due to storage etc.) for aromatic saturation. To meet the gasoline specification for the Benzene content or mixed stream where some of the Benzene is to be only saturated, than one will use a low pressure splitter to take a side draw Benzene concentrate or taken overhead and feed it to the reactor for benzene conversion to cyclohexane. If there are other lighter components in the feed, this provides an optimum condition for separation and hydrogenation rather than doing both unit operations in one equipment practiced by Reactive distillation. Reactive distillation provides non optimum conditions and requires high temperature boiling medium for the column reboiler, and on the whole the economics is not that good due to obvious reasons of catalyst cost and also process is not flexible, has to use single source packaged catalyst. The configuration is shown in FIGS. 3,4 and 5 are applicable for Benzene saturation and selective hydrogenation application as well.
Hydro-desulfurization of Gasoline/Kerosine and Diesel/Distillate:
RHT technology has differentiated itself from the existing conventional Fixed bed technologies by having the reactor installed as a side reactor so that the capital cost can be reduced by eliminating some of the equipment from the configuration and selecting multiple catalyst to enhance the catalyst productivity, optimum operating conditions and best utilization of catalyst. RHT has reduced the operating pressures compared to conventional technologies, which has a direct effect on the capital cost reduction. RHT has also removed major drawbacks of the reactive distillation technology, by operating the reaction and distillation at their respective optimum conditions. Reactive distillation has to increase the operating pressure for LCN recovery, which is not good for distillation, as the fractionation is not done at the optimum conditions. In HCN desulfurization Reactive distillation has taken most of the product overhead (rather than increasing the pressure) as mentioned in licenser's patent (U.S. Pat. No. 6,495,030), so as to increase the WABT of the catalyst zone required for HDS. Apart from having proprietary catalyst, taking most of the product overhead, one also needs to provide reflux for reactive distillation that increases the energy for this application by a factor. If one takes most of the product overhead in gasoline HDS application, the configuration becomes complex, energy costs are doubled which one can mitigate by heat integration to some degree, but unit needs much attention in operation compared to simple Fixed bed unit or RHT configuration which is a variation of Fixed bed configuration. These things are quite obvious and as regards to catalyst life in most of the operations can be the same if the reactor is designed in two phase operation or the feed has been selectively hydrogenated in the upstream equipment as shown in
In general as mentioned earlier the Fixed bed reactors would operate in single or two phase operation with certain amount of vapor at inlet or outlet as required by the design requirements. The reactors are designed to operate in upflow or downflow mode in Fixed bed operation mode.
The
The application shown in
The art of RHT process provides the best conditions for separation and reaction, and are decoupled for pressure and temperature so as meet the optimum conditions for the each process and still tied to the column so as to reduce the equipment piece count reducing the capital cost and providing the optimum yield and selectivity with standard low cost catalyst rather than expensive and catalyst, which requires cumbersome loading requirements. In the case of Benzene and LCN, fractionation can be operated at lowest possible pressure which helps the separation efficiency and also needs low level heat for reboiler, which is big saving and selectivity and yield are much better than the other processes available in market. The
RHT has developed a HDS process,
The major claim apart from improvements in reactor design and using multiple catalyst for desulfurization of the feed are as follows: RHT technology and art of the process claims that feed could be pump around or side strippers/Side draws hot liquid and desulfurizing those separately in staged reactor system where hydrogen is cascaded. This scheme provides major advantages in heat/energy savings and essentially perform the desulfurization at different severity depended on the feeds characteristics.
The disclosed method and art can be understood by referring to the attached
RHT is an emerging technology developer with a new and unique configuration, optimizing the conventional technology schemes, by reducing the equipment count by integrating the external reactor with the up stream column, this eliminates the need for stripper/stabilizer and associated equipment. This configuration allows reactors to be operated at optimum conditions, which enhances the selectivity. In isomerization of alpha to beta olefins or visa versa, and selective hydrogenation applications it provides optimum flow scheme with successive reactors to enhance the selectivity and can meet the product specifications to 10 wppm of diolefins. It provides similar approach for crude C4's from steam cracker with high content of diolefins to meet the low diolefin product as per the
RHT has developed an optimum configuration for MAPD. Acetylene, C3 to LCN hydrocarbon stream, selective hydrogenation under optimum configuration with one or multiple side reactors so as to get the best selectivity. This configuration is also good for alpha to beta isomerization of C4 to C7 olefins, especially C4 olefins. The reactors operate at 100 to 450 psig and at temperature of 100 to 500 F., and optimized for each feed and service required by the process. This provides much better selectivity than conventional or other processes like Reactive distillation which keep the products at equilibrium in the catalyst zone in isomerization reactions and if the close boilers are to isomerized, lot of catalyst will be required.
The configuration for HDS of straight run Naphtha, Kerosene and distillate is provided with classical multiple catalyst system approach so to perform the denitrification, desulfurization of the feed. The unique art is suggested here in taking the feed to HDS unit directly from the pump- arounds in a Crude/Vacuum distillation columns, the side strippers, Main Fractionator side draws, or Fractionators in the thermal crackers units, so as to save energy. This provides major savings in the Furnace as the feed is already close to the HDS reactor temperature. By using this technique provides major advantage in having the reactors operate at the desired partial pressures and the hydrogen can be staged from high pressure reactor to lower pressure HDS reactor. No doubt there could be equipment piece count has increased by the having HDS reactor for different feeds but as they are at different pressure and cascading hydrogen there are some savings in Capital and operating cost. The HDS reactor depending on the Feed characteristics could be 450 to 900 psig for the Diesel but would need higher pressure for the VGO with a high cut point Feed and high sulfur content and its refractoriness. This might need 1500 to 2000 psig or higher pressures for residue feeds, and temperatures of 700 to 750 F. and hydrogen requirements for the unit could be 600 scf/bbl to 12000 scf/bbl. So depending on the feed. Based on the feed characteristics one can always cascade the hydrogen to the lower pressure reactors. The WHSV of 0.5 to 4.0 is expected in these applications.
Other embodiments of the invention will be apparent to those skilled in the art of hydro-processing from the consideration of this specification mentioned above or from the practice of invention disclosed herein. It is intended that these specifications mentioned above be considered as exemplary only with the true scope of this invention being indicated by the following claims:
Claims
1. A process for selective hydrogenation of MAPD, acetylenes, C3 to C7 hydrocarbons, Naphtha, and for removing the mercaptan and lighter sulfur compounds in the Feed in Naphtha/gasoline range feed stocks. In doing so the dioefins in Heavy Naphtha are also saturated to olefins hence reducing the HDS reactor fouling and increasing the catalyst life and cycle lengths. The process is designed to provide mild HDS for LCN stream so as to remove the mercaptan and lighter sulfur (thiophenes) with very little olefin loss and which improves the overall octane loss on the Naphtha by reducing the olefin loss from the MCN/HCN stream. The process is also capable of saturating Benzene in reformate with high selectivity in this configuration. This configuration also provides the alpha to beta isomerization much more efficiently than any other process available in the market. The process with simultaneous reaction and distillation is not that efficient for hydro-isomerization of C4 olefins, as explained above. Especially for compounds with very small difference in boiling points, as the component takes much more stages to separate and the equilibrium is achieved in short time and the catalyst requirement for the application in reactive distillation have to be very high and are against the thermodynamic and equilibrium composition principles. But due to bulk catalyst and optimum configuration of multi staging the isomerization reaction after separation, RHT process and art stated here provides best selectivity at lowest cost. Extra equipment cost is paid off in 3 to 6 months in comparison to high catalyst/packing and installation costs. Additionally, apart from high Catalyst cost for reactive distillation, the column size is much bigger than RHT requirements. The process which use channels also are not user friendly for revamp and also the initial cost of the equipment with channels is more or less equivalent to the savings. Additionally complexity is increased as regards to processes with complex internals that have to be frequently changed. These configurations are especially shown in FIGS. 3, 4, and 5. SELECTIVE Hydrogenation Vinyl acetylene Ni, Ni/W, Ni/Mo, Pd/Ag, Pd/Pt, Pt/Zeolite etc. MAPD Pd, Pd/Ag., Ni, Ni/Mo, Ni/W, Pd/Pt, Pt/Zeolite etc. C4/C5/C6/C7 and LCN or FRCN, Ni, Pd, Ni/Mo, Ni/W, Co/Mo, Diolefin/Sulfur removal application: Pd/Pt, Pt/Zeolite etc Isomerization; Alpha to Beta Isomerization C4/C5/C6 and LCN Pd, Pd/Pt, Pd/Ag, Ni, Ni/Mo and Ni/W. Aromatic Hydrogenation Benzene Saturation Ni, Pt on Zeolite etc
- a) As per art in the claim 1, the process is optimized so as to reduce the equipment count by having the side reactors connected to the Stabilizers/Strippers that provide the best selectivity at low cost and also without having cumbersome catalyst system.
- b) As per art in the claim 1, following catalyst are used but are not limited to these catalysts only:
- c) As per art in the claim 1, the operating conditions are fine-tuned for each application based on the Feed composition and product specifications. For selective hydrogenation and benzene saturation and LCN recovery with sulfur removal, FIG. 1, 2,3,4 and 5 are the configuration basis.
- d) As per claim 1, major art in this claim as per FIG. 1, is to selective hydrogenate the feed and convert mercaptan sulfur to heavy disulfide or by mild HDS at high space velocity remove the mercaptan and light sulfur together with hydrogenating of diolefins. This technique maximizes the recovery of LCN. This is not being practiced by any licensers at present is unique to RHT process claim. RHT process as depicted in the FIGS. 1, 6 and 7 shows these configurations.
- e) As per claim 1, the high concentration diolefins are selectively hydrogenated with selective catalysts and configuration as shown in FIG. 2 configuration.
- f) As per claim 1, C3 to C7 Feed, MAPD, acetylenes selective hydrogenation, benzene saturation, and alpha to beta or visa versa isomerization is done by side reactor configuration as shown in FIGS. 3,4 and 5, which provides the best selectivity and low cost option.
- g) As per claim 1, reactors are operated in single or two phase flow (vapor/liquid mixed phase at the inlet or outlet) mode. Most of selective hydrogenation conventional processes operate in single phase or reactive distillation mode, which is two phase flow, but flowing in opposite direction. RHT design with fixed bed reactors, flow will be in one direction only. RHT is claiming the two-phase reactor or single-phase operation in all these services as one of claim. RHT reactors are designed to operate in upflow or down flow Fixed bed operation mode depending on the process requirements.
- h) As per claim 1, the multiple reactors concept as multistage with separation and reaction both being under optimum conditions, rather than single condition, provides the best selectivity and low cost operation. A multistage reactor with separation and reaction is art of this claim as shown in FIG. 4 and 5(by installing multiple beds in the column). This provides the best economics for butene-1 to butene-2, isomerization or visa versa, MAPD and acetylene, all the C3 to LCN selective hydrogenation and benzene saturation applications.
- g) As per claim 1, the FIG. 5 provides the capability to use the bulk catalyst in the column as shown for selective hydrogenation and isomerization in the column. This technique of bulk catalyst installation in the column is being claimed in any hydrogenation/isomerization or any other chemical or refinery process. This type of technique of catalyst installation is being claimed rather than the process specific, and can be used for any reaction and distillation requirements..
2. A process and art for desulfurization of the FCC gasoline under optimum condition without energy penalties and does not suffer from any proprietary catalyst requirements. The art in this process is user friendly and provides lowest octane loss for the FCC gasoline desulfurization compared to any technology at lower capital and operating cost, and has major advantages than all the processes available in the market. The art is shown in the configurations illustrated in the FIG. 6 and 7.
- a) As per claim 1 and claim 2, RHT process provides the option as shown in FIG. 1 and FIG. 6 and 7, to maximize the recovery of LCN product together with preserving olefin content, with the unique requirement of the process as described in the patent.
- b) As per claim 2, the first fractionation column (Gasoline Splitter), downstream of the mild HDS/selective hydrogenation reactor splits the Naphtha into LCN, MCN and HCN. The bottom HCN is sent to the HDS reactor after mixing with hydrogen and heating to 550 to 675 F. temperature. The reactor contains Co/Mo catalyst and could have Ni/Mo or Ni/W catalyst as the second or third bed. The MCN is fed to second or third bed after mixing with hydrogen and heating it to 525 F. to 550 F. so as to desulfurize at milder conditions. The HCN desulfurization catalyst bed is operated at most severe conditions. If there is nitrogen in the Naphtha than first bed could be Ni/W catalyst bed so to perform the denitrification. This is shown in FIG. 6 and 7.
- c) As per art of claim 2, the HDS reactor has provision to take some or all of the vapor out in between the beds and recycle the hydrogen after removing sulfur in an amine unit. The condensed hydrocarbons can be taken as product or fed back to HDS reactor for further desulfurization. This liquid is fed to the same location from where the vapor was taken out, so as to provide additional desulfurization in the catalyst bed. It will require additional hydrogen to be added into the reactor at this location. Distributor will be provided so that one can take the vapor out and send hydrogen below the distributor. This concept is shown in FIG. 7. The process by taking the vapor out and removing the H2S removes the possibility of recombination in the last bed as the partial pressure of H2S is low at that point. The RHT process provides the art that is optimum configuration for low octane loss at lowest capital and operating cost compared to other technologies.
- d) It provides the revamp of conventional or even reactive distillation technology that could reduce the octane loss and operating cost.
3. The RHT process FIG. 8, for HDS of distillate straight run or cracked/LCO and Kerosene is
- Operated at high pressure and temperature. The process uses multiple catalyst beds and multiple type of catalyst as mentioned above (Co/Mo, Ni/Mo, Ni/W, Pt/Zeolite but not limited to) to improve the product yield and economics.
- a) As per claim 3, the major claim and art is to reduce the energy requirements and also by staggering the reactor pressures for different cuts of the feed from lighter to heavier get the best operation configuration and using multiple type of catalysts suitable for HDS of the distillate Feed. These could be Co/Mo, Ni/Mo, Ni/W or Zeolite with Platinum. This provides the HDS for these distillate streams, ring opening and saturation of aromatics to enhance the cetane number of LCO or aromatic stream.
- b) As per claim 3, the art of staggering the pressures for different cut point feeds from the pump around or side draws from crude/vacuum/thermal/FCC main fractionators provides major energy savings and cost effective utilization of hydrogen. By using the technique and art mentioned here it cuts the cost and also hydrogen can be cascaded so as to reduce the compression power/energy requirements making major saving and optimizing the process.
- c) As per claim 3, the art in the process has another claim which is to use the feed to the reactors directly from the Crude and Vacuum column pump around or side strippers, Main fractionator side draw off, and also from thermal cracker side draw offs. This feed being hot provides energy saving for the furnace. As different feeds, based on the draw of cut points need different HDS severity as these will have different amount and type of sulfur, which needs milder or severe condition for reaction system.
Type: Application
Filed: Sep 27, 2006
Publication Date: Mar 27, 2008
Patent Grant number: 7959793
Applicant: Refining Hydrocarbon Technologies LLC (RHT) (Katy, TX)
Inventor: Amarjit Singh Bakshi (Katy, TX)
Application Number: 11/527,819
International Classification: C10G 45/00 (20060101);